Process for hydrodesulfurizing or hydrogenating a hydrocarbon distillate

ABSTRACT

A process for hydrotreating (hydroprocessing) hydrocarbons and mixtures of hydrocarbons utilizing a catalytic composite of a porous carrier material, a Group VI-B metal and a Group VIII metal in which process there is effected a chemical consumption of hydrogen. A specific example of one such catalyst is a composite of alumina, a molybdenum component and a cobalt component for utilization in a hydrodesulfurization process. Other hydrocarbon hydroprocesses are directed toward the hydrogenation of aromatic nuclei, the ring-opening of cyclic hydrocarbons, hydrocracking, denitrification, hydrogenation, etc.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application is a division of my application Ser. No. 628,012, filedNov. 3, 1975, now abandoned, which is a division of my application Ser.No. 571,339 filed Apr. 24, 1975, now U.S. Pat. No. 3,956,105, which inturn is a continuation-in-part of my application Ser. No. 484,519, filedJuly 1, 1974, now U.S. Pat. No. 3,935,127, all the teachings of whichapplications are incorporated herein by specific reference thereto.

APPLICABILITY OF INVENTION

The present invention encompasses the use of a catalytic composite of aporous carrier material, a Group VI-B metal component and a Group VIIImetal component. As utilized herein, the term hydrotreating is intendedto be synonymous with the term hydroprocessing, which involves theconversion of hydrocarbons at operating conditions selected to effect achemical consumption of hydrogen. Included within the processes intendedto be encompassed by the term hydroprocessing are hydrocracking,aromatic hydrogenation, ring-opening, hydrorefining (for nitrogenremoval and olefin saturation), desulfurization (often included inhydrorefining) and hydrogenation, etc. As will be recognized, one commonattribute of the processes, and the reactions being effected therein, isthat they are all hydrogen consuming, and are therefore exothermic innature.

The individual characteristics of the foregoing hydrotreating processes,including preferred operating conditions and techniques, will behereinafter described in greater detail. The subject of the presentinvention is the use of a catalytic composite which has exceptionalactivity and resistance to deactivation when employed in a hydrogenconsuming process. Such processes require a catalyst having both ahydrogenation function and a cracking function. More specifically, thepresent process uses a dual-function catalytic composite which enablessubstantial improvements in those hydroprocesses which havetraditionally used a dual-function catalyst. The particular catalyticcomposite constitutes a porous carrier material, a Group VI-B component,and a Group VIII component: specifically, an improved hydrocrackingprocess utilizes a crystalline aluminosilicate carrier material, amolybdenum compnent, and a cobalt component for improved activity,product selectivity and operation stability characteristics.

Composites having dual-function catalytic activity are widely employedin many industries for the purpose of accelerating a wide spectrum ofhydrocarbon conversion reactions. Generally, the cracking function isthought to be associated with an acid-acting material of porous,adsorptive refractory inorganic oxide type which is typically utilizedas the carrier material for a metallic component from the metals, orcompound metals, of Group V through VIII of the Periodic Table, to whichthe hydrogenation function is generally attributed.

Catalytic composites are used to promote a wide variety of hydrocarbonconversion reactions such as hydrocracking, isomerization,dehydrogenation, hydrogenation, desulfurization, reforming,ring-opening, cyclization, aromatization, alkylation andtransalkylation, polymerization, cracking, etc., some of which reactionsare hydrogen producing while others are hydrogen consuming. In using theterm hydrogen consuming, I intend to exclude those processes wherein theonly hydrogen consumption involves the saturation of light olefins,resulting from undesirable cracking, which produces the light paraffins,methane, ethane and propane. It is to the latter group of reactions,hydrogen consuming, that the present invention in applicable. In manyinstances, the commercial application of these catalysts is in processeswhere more than one of these reactions proceeds simultaneously. Anexample of this type of process is a hydrocracking process whereincatalysts are utilized to effect selective hydrogenation and cracking ofhigh molecular weight materials to produce a lower boiling, morevaluable output stream. Another such example would be the conversion ofaromatic hydrocarbons into jet fuel components, principally straight, orslightly branched paraffins.

Regardless of the reaction involved, or the particular process, it isvery important that the catalyst exhibit not only the capability toperform its specified function initially, but also perform themsatisfactorily for prolonged periods of time. The analytical termemployed in the art to measure how efficient a particular catalystperforms its intended function in a particular hydrocarbon conversionprocess, are activity, selectively and stability. For the purpose ofdiscussion, these terms are conveniently defined herein, for a givencharge stock, as follows: (1) activity is a measure of the ability ofthe catalyst to convert a hydrocarbon feedstock into products at aspecified severity level, where severity level alludes to the operatingconditions employed--the temperature, pressure, liquid hourly spacevelocity and hydrogen concentration; (2) selectivity refers to theweight percent of volume percent of the reactants that are convertedinto the desired product and/or products; (3) stability connotes therate of change of the activity and selectivity parameters withtime--obviously, the smaller rate implying the more stable catalyst.With respect to a hydrogen consuming process, for example,hydrocracking, activity, stability and selectivity are similarlydefined. Thus, "activity" connotes the quantity of charge stock, boilingabove a given temperature, which is converted to hydrocarbons boilingbelow the given temperature. "Selectivity" refers to the quantity ofconverted charge stock which boils below the desired end point of theproduct, as well as above a minimum specified initial boiling point."Stability" connotes the rate of change of activity and selectivity.Thus, for example, where gas oil, boiling above 650° F., is subjected tohydrocracking, "activity" connotes the converstion of 650° F.-pluscharge stock to 650° F-minus product. "Selectivity" can allude to thequantity of conversion into gasoline boiling range hydrocarbons--i.e.,pentanes, and heavier, normally liquid hydrocarbons boiling up to about400° F. "Stability" might be conveniently expressed in terms oftemperature increase required during various increments of catalystlife, in order to maintain a desired activity.

As is well known to those skilled in the art, the principal cause ofobserved deactivation or instability of a duel-function catalyst isassociated with the fact that coke forms on the surface of the catalystduring the course of the reaction. More specifically, in the varioushydrocarbon conversion processes, and especially those which arecategorized as hydrogen consuming, the operating conditions utilizedresult in the formation of high molecular weight, black, solid orsemi-solid, hydrogen-poor carbonaceous material which coats the surfaceof the catalyst and reduces its activity by shielding its active sitesfrom the reactants. Accordingly, a major problem facing workers in thisarea is the development of more active and selective catalyticcomposites that are not as sensitive to the presence of thesecarbonaceous materials and/or have the capability to suppress the rateof formation of these materials at the operating conditions employed ina particular process.

I have now found a dual-function catalytic composite which possessesimproved activity, selectivity and stability when employed in thehydroprocessing of hydrocarbons, wherein there is effected a chemicalconsumption of hydrogen. In particular, I have found that the use of acatalytic composite of a Group VI-B metallic component, and a Group VIIImetallic component with a porous carrier material which is prepared by(a) dry mixing a finely divided Group VI-B metal compound, a Group VIIImetal compound and a refractory inorganic oxide, said metal compoundbeing utilized in an amount to provide from about 25 to about 90% of theGroup VI-B metal component of the final catalytic composite, and fromabout 25 to about 90% of the group VIII metal component; (b) peptizingthe mixture and forming an extrudable dough; (c) extruding said doughand calcining the extrudate; (d) impragnating the calcined extrudatewith a Group VI-B metal compound and a Group VIII metal compound, saidmetal compound being utilized in an amount to provide a final catalyticcomposite containing from about 4 to about 30 weight percent Group VI-Bmetal and from about 0.1 to about 10 weight percent Group VIII metal;and, (e) calcining the resulting composite in an oxidizing atmosphere,improves the overall operation of these hydrogen-consuming processes.Moreover, I have determined that a catalytic composite of a crystallinealuminosilicate carrier material, a Group VI-B metallic component and aGroup VIII metallic component prepared as described hereinabove, whenutilized in a process for hydrocracking hydrocarbonaceous material intolower-boiling hydrocarbon products, affords substantial improvement inperformance and results. As indicated, the present invention essentiallyinvolves the use of a catalyst which comprises a Group VI-B metalliccomponent, a Group VIII metallic component and a porous carrier materialprepared as described hereinabove, and enables the performancecharacteristics of the process to be sharply and materially improved.

This invention also relates to the hydroprocessing of petroleumhydrocarbon fractions such as residual fuel oils, and to a method forthe manufacture of a catalytic composite particularly adapted thereto.It has become well known that oxides of sulfur, plus lesser amounts ofother sulfurous compounds, are among the major pollutants of theatmosphere. It has been estimated that, in this country alone, in excessof about 23 million tons of sulfur dioxide has been discharged into theatmosphere on an annual basis. The increasingly deleterious effect ofthe sulfurous pollutants with respect to cardiorespiratory disease, eyeirritation, and the like, has prompted rather severe legislature actionto control the amount of sulfur dioxide discharged into the atmosphere,particularly in densely populated areas where the problem is more acute.It has been recognized that the combustion of petroleum productsaccounts for a substantial portion of said oxides of sulfur andlegislation has been effected or proposed which is particularly directedto the limitation of sulfurous compounds in residual fuel oils to beburned in densely populated areas. The supply of residual fuel oils ofsuitably low sulfur content is entirely inadequate for the present dayrequirements and it becomes increasingly important to develop improveddesulfurization techniques to treat the more accessible and abundantresidual fuel oils of relatively high sulfur content.

Desulfurization technology is presently concerned with hydrotreating andto the development of catalysts that are more selective and/or operateat less severe conditions to obviate hydrocracking of the residual fueloil. Hydrotreating, or hydrodesulfurization, is generally effected athydrodesulfurization reaction conditions including an imposed hydrogenpressure of from about 100 to about 3000 psi. Normally, the hydrogen ischarged together with recycle hydrogen to provide from about 1000 toabout 50,000 standard cubic feet per barrel of hydrocarbon charge.Hydrodesulfurization reaction conditions further include an elevatedtemperature, usually from about 200° to about 800° F. althoughtemperatures in the higher range, say from about 600° to about 800° F.,are most suitable. Also, a sulfurcontaining feed stock is generallysuitably processed at a liquid hourly space velocity of from about 0.5to about 20. Hydrodesulfurization catalysts preferably comprise a GroupVI-B metal, usually molybdenum, and a Group VIII metal, usually nickelor cobalt, on a refractory inorganic oxide carrier material, usuallyalumina.

OBJECTS AND EMBODIMENT

It is an object of this invention to present a novel method for themanufacture of a catalytic composite, of a Group VI-B metal, a GroupVIII metal and a refractory inorganic oxide carrier material, the methodbeing particularly applicable to the manufacture of an improvedhydrorefining catalyst.

In one of its broad aspects, the present invention embodies a method ofmanufacturing a catalytic composite of from about 4 to about 30 weightpercent Group VI-B metal, from about 0.1 to about 10 weight percentGroup VIII metal and a refractory inorganic oxide carrier material whichcomprises dry mixing a finely divided Group VI-B metal compound, a GroupVIII metal compound and a refractory inoganic oxide, said metalcompounds being utilized in an amount to provide from about 25 to about90% of the Group VI-B metal component of the final catalytic composite,and from about 25 to about 90% of the Group VIII metal component;peptizing the mixture and forming an extrudable dough; extruding saiddough, and calcining the extrudate; impregnating the calcined extrudatewith a Group VI-B metal compound and a Group VIII metal compound, saidmetal compound being utilized in an amount to provide a final catalyticcomposite containing from about 4 to about 30 weight percent Group VI-Bmetal and from about 0.1 to about 10 weight percent Group VIII metal;and calcining the resulting composite in an oxidizing atmosphere.

One of the more specific embodiments of the present invention relates toa method of manufacturing a catalytic composite of from about 4 to about30 weight percent molybdenum, from about 0.1 to about 10 weight percentcobalt and alpha-alumina monohydrate, which comprises dry mixing finelydivided molybdic anhydride, cobalt carbonate and alpha-aluminamonohydrate, said molybdic anhydride and said cobalt carbonate beingutilized in an amount to provide from about 25 to about 90% of themolybdenum component and from about 25 to about 90% of the cobaltcomponent; peptizing the mixture with dilute nitric acid and forming anextrudable dough; extruding said dough, and calcining the extrudate at atemperature of from about 600° to about 1200° F.; impregnating thecalcined extrudate with a common ammoniacal solution of molybdic acidand cobalt nitrate, said molybdic acid and said cobalt nitrate beingutilized in an amount to provide a final catalytic composite containingfrom 4 to about 30 weight percent molybdenum and from about 1 to about10 weight percent cobalt; calcining the resulting composite in air at atemperature of from about 600° to about 1200° F.

Another object of the present invention is to afford a process for thehydroprocessing of a hydrocarbon, or mixtures of hydrocarbons. Acorollary objective is to improve the selectivity and stability ofhydroprocessing utilizing a highly active, Group VI-Bcomponent-containing and a Group VIII component-containing catalyticcomposite.

A specific object of my invention resides in the improvement ofhydrogen-consuming processes including hydrocracking, hydrorefining,ring-opening for jet fuel production, hydrogenation of aromatichydrocarbons, desulfurization, denitrification, etc. Therefore, in oneembodiment, the present invention encompasses a hydrocarbon hydroprocesswhich comprises reacting a hydrocarbon with hydrogen at conditionsselected to effect chemical consumption of hydrogen and in contact witha catalytic composite of a Group VI-B component, a Group VIII componentand a porour carrier material.

In another embodiment, the operating conditions include a pressure offrom 400 to about 5,000 psig., an LHSV (defined as volumes of liquidhydrocarbon charge per hour per volume of catalyst disposed in thereaction zone) of from 0.1 to about 10.0, a hydrogen circulation rate offrom 1,000 to about 50,000 scf./Bbl. and a maximum catalyst temperatureof from 200° to about 900° F.

In another embodiment, the process is further characterized in that thecatalytic composite is reduced and sulfided prior to contacting thehydrocarbon feed stream. In still another embodiment, my inventioninvolves a process for hydrogenating a cokeforming hydrocarbondistillate containing di-olefinic and mono-olefinic hydrocarbons, andaromatics, which process comprises reacting said distillate withhydrogen, at a temperature below about 500° F., in contact with acatalytic composite of an alumina-containing refractory inorganic oxide,a Group VIII component, an alkali metal component, and a Group VI-Bcomponent, and recovering an aromatic/mono-olefinic hydrocarbonconcentrate substantially free from conjugated di-olefinic hydrocarbons.

Another embodiment affords a catalytic composite comprising asubstantially pure crystalline aluminosilicate material, at least about90.0% by weight of which is zeolitic, a Group VIII component, and aGroup VI-B component.

Other objects and embodiments of my invention relate to additionaldetails regarding preferred catalytic ingredients, the concentration ofcomponents in the catalytic composite, methods of catalyst preparation,individual operating conditions for use in the various hydrotreatingprocesses, preferred processing techniques and the like particularswhich are hereinafter given in the following, more detailed summary ofmy invention.

SUMMARY OF THE INVENTION

As hereinabove set forth, the present invention involves thehydroprocessing of hydrocarbons and mixtures of hydrocarbons, utilizinga particular catalytic composite. This catalyst comprises a porouscarrier material having combined therewith a Group VIII metalliccomponent, and a Group VI-B metallic component; in many applications,the catalytic composite will also contain a halogen component, and insome select application, an alkali metal or alkaline-earth metalcomponent. Considering first the porous carrier material, it ispreferred that it be a porous, adsorptive, high-surface area supporthaving a surface area of about 25 to about 500 square meters per gram.The porous carrier material is necessarily relatively refractory withrespect to the operating conditions employed in the particularhydrotreating process, and it is intended to include carrier materialswhich have traditionally been utilized in dual-function hydrocarbonconversion catalyst. In particular, suitable carrier materials areselected from the group of amorphous refractory inorganic oxidesincluding alumina, titania, zirconia, chromia, magnesia, thoria, boria,silica-alumina, silica-magnesia, chromia-alumina, alumina-boria,alumina-silica-boron phosphate, silica-zironia, etc. When of theamorphous type, the preferred carrier material is a composite of aluminaand silica with silica being present in an amount of about 10.0% toabout 90.0% by weight.

In many hydroprocessing applications of the present invention,particularly hydrocracking heavy hydrocarbonaceous material to producelower-boiling hydrocarbon products, the carrier material will constitutea crystalline aluminosilicate, often referred to as being zeolitic innature. This may be naturally occurring, or synthetically prepared, andincludes mordenite, faujasite, Type A or Type U molecular sieves, etc.When utilized as the carrier material, the zeolitic material may be inthe hydrogen form, or in a form which has been treated with multi-valentcations.

As hereinabove set forth, the porous carrier material, for use in theprocess of the present invention, is a refractory inorganic oxide,either alumina in and of itself, or in combination with one or moreother refractory inorganic oxides, and particularly in combination withsilica. When utilized as the sole component of the carrier material, thealumina may be of the gamma-, eta, or theta-alumina type, with gamma-,or eta-alumina giving the best results. In addition, the preferredcarrier materials have an apparent bulk density of about 0.30 to about0.70 gm/cc. and surface area characteristics such that the average porediameter is about 20 to about 300 Angstroms, the pore volume is about0.10 to about 1.0 milliliters per gram and the surface area is about 100to about 500 square meters per gram. Whatever type of refractoryinorganic oxide is employed, it may be activated prior to use by one ormore treatments including drying, calcination, streaming, etc.

When a crystalline aluminosilicate, or zeolitic material, is intendedfor use as the carrier, it may be prepared in a number of ways. Onecommon way is to mix solutions of sodium silicate, or colloidal silica,and sodium aluminate, and allow these solutions to react to form a solidcrystalline aluminosilicate. Another method is to contact a solidinorganic oxide, from the group of silica, alumina, and mixturesthereof, with an aqueous treating solution containing alkali metalcations (preferably sodium) and anions selected from the group ofhydroxyl, silicate and aluminate, and allow the solid-liquid mixture toreact until the desired crystalline aluminosilicate has been formed. Oneparticular method is especially preferred when the carrier material isintended to be a crystalline aluminosilicate. This stems from the factthat the method can produce a carrier material of substantially purecrystalline aluminosilicate particles. In employing the term"substantially pure," the intended connotation is an aggregate particleat least 90.0% by weight of which is zeolitic. Thus, this carrier isdistinguished from an amorphous carrier material, or prior art pillsand/or extrudates in which the zeolitic material might be dispersedwithin an amorphous matrix with the result that only about 40.0 to about70.0% by weight of the final particle is zeolitic. The preferred methodof preparing the carrier material produces crystalline aluminosilicatesof the faujasite modification, and utilizes aqueous solutions ofcolloidal silica and sodium aluminate. Colloidal silica is a suspensionin which the suspended particles are present in very finely divided form-- i.e., having a particle size from about 1 to about 500 millimicronsin diameter. The type of crystalline aluminosilicate which is producedis primarily dependent upon the conditions under which crystallizationoccurs, with the SiO₂ /Al₂ O₃ ratio, the Na₂ O/SiO₂ ratio, the H₂ O/Na₂O ratio, temperature and time being the important variables.

After the solid crystalline aluminosilicate has been formed, the motherliquor is separated from the solids by methods such as decantation orfiltration. The solids are water-washed and filtered to removeundesirable ions, and to reduce the quantity of amorphous material, andare then reslurried in water to a solids concentration of about 5.0% toabout 50.0%. The cake and the water are violently agitated andhomogenized until the agglomerates are broken and the solids areuniformly dispersed in what appears to be a colloidal suspension. Thesuspension is then spray dried by conventional means such as pressuringthe suspension through an orifice into a hot, dry chamber. The solidparticles are withdrawn from the drying chamber and are suitable forforming into finished extrudate particles of desired size and shape.

In accordance with the method of this invention, a finely divided GroupVI-B metal compound, Group VIII metal compound and a refractoryinorganic oxide are dry mixed, the mixture being subsequently peptizedto form an extrudable dough. The expression "finely divided" isdescriptive of particles having an average diameter of less than about150 microns, for example, particles which are recoverable through a 105micron microsieve. The refractory inorganic oxide can be alumina,silica, zirconia, thoria, boria, chromium, magnesia, titania, and thelike, or composites thereof such as alumina-silica, alumina-zirconia,and the like. Alumina is a preferred refractory inorganic oxide,especially alpha-alumina monohydrate of the boehmite structure, and thefurther description of the method of this invention is presented withrespect thereto. The dry mixing operation is improved utilizing analpha-alumina monohydrate characterized by a weight loss of ignition at900° C. of from about 20 to about 30 weight percent. In addition to itscontribution to the catalytic properties of the catalytic composite ofthis invention, the alpha-alumina monohydrate improves the extrusioncharacteristics of the mixture whereby the mixture is readily extrudedthrough a 1/32 - 1/8 inch orifice at a pressure of less than about 500psig.

Molybdic anhydride is a particularly suitable Group VI-B metal compound,and cobalt carbonate is a particularly suitable Group VIII metalcompound for dry mixing with the alpha-alumina monohydrate as hereincontemplated. Other suitable Group VI-B metal compounds, that is,compounds of molybdenum, tungsten and chromium, include molybdic acid,ammonium molybdate, ammonium chromate, chromium acetate, chromouschloride, chromium nitrate, tungstic acid, etc. Other Group VIII metalcompounds which may be employed, that is, compounds of iron, nickel,cobalt, platinum, palladium, ruthenium, rhodium, osmium and irridium,include nickel nitrate, nickel sulfate, nickel chloride, nickel acetate,cobaltous sulfate, ferric nitrate, ferric sulfate, platinum chloride,palladium chloride and the like. In any case, the resulting mixture ispeptized, shitably by the addition thereto of a weak acid such as formicacid, acetic acid, propionic acid, and the like, although the strongeracids such as sulfuric acid, hydrochloric acid, and particularly nitricacid are preferred. In the case where a crystalline aluminosilicatematerial is selected as an ingredient, a suitably weak peptizing agentwill be utilized to avoid degradation of the crystalline structure whichdegradation subsequently produces a low activity catalyst. Sufficientpeptizing agent is blended or mulled with the mixture to form anextrudable dough or pliable plastic mass.

The extrusion operation is suitably effected with commercial extrusionapparatus. For example, the dough is continuously processed through acylinder by means of a rotating screw, and pressured through aperforated plate at one end of the cylinder. The extrudate may be cutinto particles of desired length prior to drying and calcining by meansof a rotating knife as the extrudate emerges from the perforated plate.

Alternatively, the extrudate may be broken into particles of randomlength during the drying and calcining process. In any case, theextrudate is calcined, calcining being preferably effected in anoxidizing atmosphere such as air at a temperature of from about 600° toabout 1200° F. over a period of from about 2 to about 4 hours.

The catalytic composite of this invention is prepared to contain fromabout 4 to about 30 weight percent Group VI-B metal and from about 0.1to about 10 weight percent Group VIII metal. Only a fraction of thetotal desired metals content of the final catalytic composite is addedthereto by the foregoing co-extrusion technique. More particularly, saidGroup VI-B metal compound and said Group VIII metal compound areutilized in an amount to provide from about 25% to about 90% of each ofthe Group VI-B metal component and the Group VIII metal of the finalcatalytic composite. The remainder of the desired total metals contentis added by impregnating the calcined extrudate with a Group VI-B metalcompound and a Group VIII metal compound.

It is common practice to deposit catalytically active metalliccomponents on a support or carrier material by the method whereby asoluble compound of the desired metallic component is impregnated on thecarrier material from an aqueous solution. The soluble compound servesas a precursor of the metallic component such that, upon subsequentheating of the impregnated carrier material at a temperature effectingdecomposition of said compound, the desired metallic component is formedand deposited upon the carrier material. The aqueous impregnatingsolution will thus comprise a soluble precursor compound of a Group VI-Bmetal. Suitable compounds include ammonium molybdate, ammoniumparamolybdate, molybdic acid, ammonium chromate, ammoniumperoxychromate, chromium acetate, chromous chloride, chromium nitrate,ammonium metatungstate, tungstic acid, etc. The impregnating solution issuitably a common solution of a Group VI-B metal compound and a GroupVIII metal compound. Suitable soluble compounds of Group VIII metalsinclude nickel nitrate, nickel sulfate, nickel chloride, nickel bromide,nickel fluoride, nickel iodide, nickel acetate, nickel formate,cobaltous nitrate, cobaltous sulfate, cobaltous fluoride, ferricfluoride, ferric bromide, ferric nitrate, ferric sulfate, ferricformate, ferric acetate, platinum chloride, chloroplatinic acid,chloropalladic acid, palladium chloride, etc. Of the Group VI-B metals,molybdenum is preferred.

Impregnation of the calcined extrudate can be accomplished byconventional techniques whereby the extrudate particles are soaked,dipped, suspended or otherwise immersed in the impregnating solution atconditions to absorb a soluble compound comprising the desired catalyticcomponent. Certain impregnating techniques have been found to beparticularly favorable to promote desired physical properties of thefinished catalyst. Thus, impregnation of the Group VI-B and Group VIIImetal components is preferably from a common aqueous ammoniacal solutionof soluble compounds thereof, for example, an ammoniacal solution ofmolybdic acid and cobalt nitrate. Further the impregnation is preferablyeffected utilizing a minimal volume of impregnating solutioncommensurate with an even distribution of the catalytic components onthe calcined extrudate particles. One preferred method involves the useof a steam-jacketed rotary dryer. The extrudate particles are immersedin the impregnating solution contained in the dryer and tumbled thereinby the rotating motion of the dryer, the volume of extrudate particlesso treated being initially in the range of from about 0.7 to about 1.0with respect to the volume of the impregnating solution. Evaporation ofthe solution in contact with the extrudate particles is expedited byapplying steam to the dryer jacket. The evaporation is furtherfacilitated by a continuous purge of the dryer utilizing a flow of drygas, suitably air or nitrogen. The impregnated particles, thus dried,are thereafter calcined in an oxygen-containing atmosphere at atemperature of from about 600° to about 1200° F. in accordance withprior art practice, usually for a period of from about 1 to about 8hours or more.

Although not essential to successful hydroprocessing in all cases, infact detrimental in some, a halogen component may be incorporated intothe catalytic composite. Accordingly, a preferred catalytic composite,for use in the present process, comprises a combination of a cobaltcomponent, a molybdenum component and a halogen component. Although theprecise form of the chemistry of the association of the halogencomponent with the carrier material and metallic components is notaccurately known, it is customary in the art to refer to the halogencomponent as being combined with the carrier material, or with the otheringredients of the catalyst. The combined halogen may be eitherfluorine, chlorine, iodine, bromine, or mixtures thereof. Of these,fluorine and particularly chlorine are preferred for the hydrocarbonhydroprocesses encompassed by the present invention. The halogen may beadded to the catalyst in any suitable manner. For example, the halogenmay be added at any convenient stage during the preparation. Halogen maysuitably be added to the calcined extrudate, to the calcined extrudateduring metal impregnation or to the calcined impregnated extrudate. Thehalogen may be added as an aqueous solution of an acid such as hydrogenfluoride, hydrogen chloride, hydrogen bromide, hydrogen iodide, etc. Thequantity of halogen is such that the final catalytic composite containsabout 0.1% to about 1.5% by weight, and preferably from about 0.5% toabout 1.2% calculated on an elemental basis.

In embodiments of the present invention wherein the instant catalyticcomposite is used for the hydrogenation of hydrogenatable hydrocarbons,it is ordinarily a preferred practice to include an alkali or alkalineearth metal component in the composite. More precisely, this optionalcomponent is selected from the group consisting of the compounds of thealkali metals -- cesium, rubidium, potassium, sodium, and lithium -- andthe compounds of the alkaline earth metals -- calcium, strontium, bariumand magnesium. Generally, good results are obtained in these embodimentswhen this component constitutes about 1 to about 5 weight percent of thecomposite, calculated on an elemental basis. This optional alkali oralkaline earth metal component can be incorporated in the composite inany of the known ways, with impregnation with an aqueous solution of asuitable water-soluble decomposable compound being preferred.

An optional ingredient for the catalyst of the present invention is aFriedel-Crafts metal halide component. This ingredient is particularlyuseful in hydrocarbon conversion embodiments of the present inventionwherein it is preferred that the catalyst utilized has a strong acid orcracking function associated therewith -- for example, an embodimentwherein hydrocarbons are to be hydrocracked or isomerized with thecatalyst of the present invention. Suitable metal halides of theFriedel-Crafts type include aluminum chloride, aluminum bromide, ferricchloride, ferric bromide, zinc chloride and the like compounds, with thealuminum halides and particularly aluminum chloride ordinarily yieldingbest results. Generally, this optional ingredient can be incorporatedinto the composite of the present invention by any of the conventionalmethods for adding metallic halides of this type; however, best resultsare ordinarily obtained when the metallic halide is sublimed onto thesurface of the carrier material according to the preferred methoddisclosed in U.S. Pat. No. 2,999,074. The component can generally beutilized in any amount which is catalytically effective, with a valueselected from the range of about 1 to about 100 weight percent of thecarrier material generally being preferred. When used in many of thehydrogen-consuming processes hereinbefore described, the foregoingquantities of metallic components will be combined with a carriermaterial of alumina and silica, wherein the silica concentration is 10.0to about 90.0% by weight.

Regardless of the details of how the components of the catalyst arecombined with the porous carrier material, the final catalyst generallywill be calcined or oxidized at a temperature of about 700° F. to about1100° F. in an air atmosphere for a period of about 0.5 to about 10hours in order to convert substantially all of the metallic componentssubstantially to the oxide form. Because a halogen component may beutilized in the catalyst, best results are generally obtained when thehalogen content of the catalyst is adjusted during the calcination stepby including a halogen or a halogen-containing compound in the airatmosphere utilized. In particular, when the halogen component of thecatalyst is chlorine, it is preferred to use a mole ratio of H₂ O to HClof about 5:1 to about 100:1 during at least a portion of the calcinationstep in order to adjust the final chlorine content of the catalyst to arange of about 0.5 to about 1.5 weight percent.

The resulting catalytic composite may, in some cases, be beneficiallysubjected to a presulfiding operating designed to incorporate in thecatalytic composite from about 0.05 to about 0.5 weight percent sulfurcalculated on an elemental basis. Preferably, this presulfidingtreatment takes place in the presence of hydrogen and a suitablesulfur-containing compound such as hydrogen sulfide, lower molecularweight mercaptans, organic sulfides, etc. Typically, this procedurecomprises treating the catalyst with a sulfiding gas such as a mixtureof hydrogen and hydrogen sulfide having about 10 moles of hydrogen permole of hydrogen sulfide at conditions sufficient to effect the desiredincorporation of sulfur, generally including a temperature ranging fromabout 50° up to about 1100° F. or more. It is generally a good practiceto perform this presulfiding step under substantially water-freeconditions.

According to the present invention, a hydrocarbon charge stock andhydrogen are contacted with a catalyst of the type described above in ahydrocarbon conversion zone. This contacting may be accomplished byusing the catalyst in a fixed bed system, a moving bed system, afluidized bed system, or in a batch type operation; however, in view ofthe danger of attrition losses of the valuable catalyst and of wellknown operational advantages, it is preferred to use a fixed bed system.In this system, a hydrogen-rich gas and the charge stock are preheatedby any suitable heating means to the desired reaction temperature andthen are passed, into a conversion zone containing a fixed bed of thecatalyst type previously characterized. It is, of course, understoodthat the conversion zone may be one or more separate reactors withsuitable means therebetween to insure that the desired conversiontemperature is maintained at the entrance to each reactor. It is alsoimportant to note that the reactants may be contacted with the catalystbed in either upward, downward, or radial flow fashion with the latterbeing preferred. In addition, the reactants may be in the liquid phase,a mixed liquid-vapor phase, or a vapor phase when they contact thecatalyst.

The operating conditions imposed upon the reaction zones are dependentupon the particular hydroprocessing being effected. However, theseoperating conditions will include a pressure from about 400 to about5,000 psig., a liquid hourly space velocity of about 0.1 to about 10.0and a hydrogen circulation rate within the range of about 1,000 to about50,000 standard cubic feet per barrel. In view of the fact that thereactions being effected are exothermic in nature, an increasingtemperature gradient is experienced as the hydrogen and feed stocktraverse the catalyst bed. For any given hydrogen-consuming process, itis desirable to maintain the maximum catalyst bed temperature belowabout 900° F., which temperature is virtually identical to thatconveniently measured at the outlet of the reaction zone.Hydrogen-consuming processes are conducted at a temperature in the rangeof about 200° to about 900° F., and it is intended herein that thestated temperature of operation alludes to the maximum catalyst bedtemperature. In order to assure that the catalyst bed temperature doesnot exceed the maximum allowed for a given process, the use ofconventional quench streams, either normally liquid or gaseous,introduced at one or more intermediate loci of the catalyst bed, may beutilized. In some of the hydrocarbon hydroprocesses encompassed by thepresent invention, and especially where hydrocracking a heavyhydrocarbonaceous material to produce lower-boiling hydrocarbonproducts, that portion of the normally liquid product effluent boilingabove the end point of the desired product will be recycled to combinewith the fresh hydrocarbon charge stock. In these situations, thecombined liquid feed ratio (defined as volume of total liquid charge tothe reaction zone per volume of fresh feed charge to the reaction zone)will be within the range of about 1.1 to about 6.0.

Specific operating conditions, processing techniques, particularcatalytic composites and other individual process details will be givenin the following detailed description of several of the hydrocarbonhydroprocesses to which the present invention is applicable. Thefollowing examples are presented in illustration of my invention. Inpresenting these examples, it is not intended that the invention belimited to the specific illustrations, nor is it intended that a givenprocess be limited to the particular operating conditions, catalyticcomposite, processing techniques, charge stock, etc. It is understood,therefore, that the present invention is merely illustrated by thespecifics hereinafter set forth.

EXAMPLE I

In this example, the present invention is illustrated as applied to thehydrogenation of aromatic hydrocarbons such as benzene, toluene, thevarious xylenes, naphthalenes, etc., to form the corresponding cyclicparaffins. When applied to the hydrogenation of aromatic hydrocarbons,which are contaminated by sulfurous compounds, primarily thiopheniccompounds, the process is advantageous in that it affords 100.0%conversion without the necessity for the substantially complete priorremoval of the sulfur compounds. The corresponding cyclic paraffins,resulting from the hydrogenation of the aromatic nuclei, includecompounds such as cyclohexane, mono-, di-, tri-substituted cyclohexanes,decahydronaphthalene, tetrahydronaphthalene, etc., which find widespreaduse in a variety of commercial industries in the manufacture of nylon,as solvents for various fats, oils, waxes, etc.

Aromatic concentrates are obtained by a multiplicity of techniques. Forexample, a benzene-containing fraction may be subjected to distillationto provide a heart-cut which contains the benzene. This is thensubjected to a solvent extraction process which separates the benzenefrom the normal or iso-paraffinic components, and the naphthenescontained therein. Benzene is readily recovered from the selectedsolvent by way of distillation, and in a purity of 99.0% or more. Inaccordance with the present process, the benzene is hydrogenated incontact with a catalytic composite containing about 4 to about 30 weightpercent Group VI-B metal, from about 0.1 to about 30 weight percentGroup VIII metal, and from about 0.01% to about 1.5% by weight of analkalinous metal component. Operating conditions include a maximumcatalyst bed temperature in the range of about 200° to about 800° F., apressure of from 500 to about 2,000 psig., a liquid hourly spacevelocity of about 1.0 to about 10.0 and a hydrogen circulation rate inan amount sufficient to yield a mole ratio of hydrogen to cyclohexane,in the product effluent from the last reaction zone, not substantiallyless than about 4.0:1. Although not essential, one preferred operatingtechnique involves the use of three reaction zones, each of whichcontains approximately one-third of the total quantity of catalystemployed. The process is further facilitated when the total freshbenzene is added in three approximately equal portions, one each to theinlet of each of the three reaction zones.

The catalyst utilized is a substantially halogen-free alumina carriermaterial combined with about 3.2 weight percent cobalt, 8.7 weightpercent molybdenum, and about 0.90% by weight of lithium, all of whichare calculated on the basis of the elemental metals. The hydrogenationprocess will be described in connection with a commercially-scaled unithaving a total fresh benzene feed capacity of about 1,488 barrels perday. Make-up gas in an amount of about 741.6 mols/hr. is admixed with2,396 Bbl./day (about 329 mols/hr.) of a cyclohexane recycle stream, themixture being at a temperature of about 137° F., and further mixed with96.24 mols/hr. (582 Bbl./day) of the benzene feed: the final mixtureconstitutes the total charge to the first reaction zone.

Following suitable heat-exchange with various hot effluent streams, thetotal feed to the first reaction zone is at a temperature of 385° F. anda pressure of 460 psig. The reaction zone effluent is at a temperatureof 606° F. and a pressure of about 450 psig. The total effluent from thefirst reaction zone is utilized as a heat-exchange medium, in a steamgenerator, whereby the temperature is reduced to a level of about 545°F. The cooled effluent is admixed with about 98.5 moles per hour (596Bbl./day) of fresh benzene feed, at a temperature of 100° F.; theresulting temperature is 400° F., and the mixture enters the secondreaction zone at a pressure of about 440 psig. The second reaction zoneeffluent, at a pressure of 425 psig. and a temperature of 611° F., isadmixed with 51.21 mols/hr. (310 Bbl./day) of fresh benzene feed, theresulting mixture being at a temperature of 578° F. Following its use asa heat-exchange medium, the temperature is reduced to 400° F., and themixture enters the third rection zone at a pressure of 415 psig. Thethird reaction zone effluent is at a temperature of about 509° F. and apressure of about 400 psig. Through utilization as a heat-exchangemedium, the temperature is reduced to a level of about 244° F., andsubsequently reduced to a level of about 115° F. by use of an air-cooledcondenser. The cooled third reaction zone effluent is introduced into ahigh pressure separator, at a pressure of about 370 psig.

A hydrogen-rich vaporous phase is withdrawn from the high pressureseparator and recycled by way of compressive means, at a pressure ofabout 475 psig., to the inlet of the first reaction zone. A portion ofthe normally liquid phase is recycled to the first reaction zone as thecyclohexane concentrate hereinbefore described. The remainder of thenormally liquid phase is passed into a stabilizing column functioning atan operating pressure of about 250 psig., a top temperature of about160° F. and a bottom temperature of about 430° F. The cyclohexaneproduct is withdrawn from the stabilizer as a bottoms stream, theoverhead stream being vented to fuel. The cyclohexane concentrate isrecovered in an amount of about 245.80 moles per hour, of which onlyabout 0.60 moles per hour constitutes other hexanes. In brief summation,of the 19,207 pounds per hour of fresh benzene feed, 20,685 pounds perhour of cyclohexane product is recovered.

EXAMPLE II

Another hydrocarbon hydroprocessing scheme, to which the presentinvention is applicable, involves the hydrorefining of coke-forminghydrocarbon distillates. These hydrocarbon distillates are generallysulfurous in nature, and contain mono-olefinic, di-olefinic and aromatichydrocarbons. Through the utilization of a catalytic composite preparedaccording to the present invention, increased selectivity and stabilityof operation is obtained; selectivity is most noticeable with respect tothe retention of aromatics, and in hydrogenating conjugated di-olefinicand mono-olefinic hydrocarbons. Such charge stocks generally result indiverse conversion processes including the catalytic and/or thermalcracking of petroleum, sometimes referred to as pyrolysis, thedestructive distillation of wood or coal, shale oil retorting, etc. Theimpurities in these distillate fractions must necessarily be removedbefore the distillates are suitable for their intended use, or whichwhen removed, enhance the value of the distillate fraction for furtherprocessing. Frequently, it is intended that these charge stocks besubstantially desulfurized, saturated to the extent necessary to removethe conjugated di-olefins, while simultaneously retaining the aromatichydrocarbons. When subjected to hydrorefining for the purpose ofremoving the contaminating influences, there is encountered difficultyin effecting the desired degree of reaction due to the formation of cokeand other carbonaceous material.

As utilized herein, "hydrogenating" is intended to be synonymous with"hydrorefining." The purpose is to provide a highly selective and stableprocess for hydrogenating coke-forming hydrocarbon distillates, and thisis accomplished through the use of a fixed-bed catalytic reaction systemutilizing a catalyst prepared according to the present invention. Thereexists two separate, desirable routes for the treatment of coke-formingdistillates, for example a pyrolysis naphtha by-product. One such routeis directed toward a product suitable for use in certain gasolineblending. With this as the desired object, the process can be effectedin a single stage, or reaction zone, with the catalytic compositehereinafter specifically described as the first-stage catalyst. Theattainable selectivity in this instance resides primarily in thehydrogenation of highly reactive double bonds. In the case of conjugateddi-olefins, the selectivity afforded restricts the hydrogenation toproduce mono-olefins, and, with respect to the styrenes, for example,the hydrogenation is inhibited to produce alkyl benzenes without "ring"saturation. The selectivity is accomplished with a minimum of polymerformation either to "gums," or lower molecular weight polymers whichwould necessitate a re-running of the product before blending togasoline would be feasible. Other advantages of restricting thehydrogenating of the conjugated di-olefins, such as 1,5 normal hexadieneare not usually offensive in suitably inhibited gasolines in somelocales, and will not react in this first stage. Some fresh chargestocks are sufficiently low in mercaptan sulfur content that directgasoline blending may be considered, although a mild treatment formercaptan sulfur removal might be necessary. These considerations aregenerally applicable to foreign markets, particularly European, whereolefinic and sulfur-containing gasolines are not too objectionable. Itmust be noted that the sulfurous compounds, and the mono-olefins,whether virgin, or products of di-olefin partial saturation, areunchanged in the single, or first-stage reaction zone. Where however thedesired end result is aromatic hydrocarbon retention, intended forsubsequent extraction, the two-stage route is required. The mono-olefinsmust be substantially saturated in the second stage to facilitatearomatic extraction by way of currently utilized methods. Thus, thedesired necessary hydrogenation involves saturation of the mono-olefins,as well as sulfur removal, the latter required for an acceptableultimate aromatic product. Attendant upon this is the necessity to avoideven partial saturation of aromatic nuclei.

With respect to one catalytic zone, its principal function involves theselective hydrogenation of conjugated di-olefinic hydrocarbons tomono-olefinic hydrocarbons. This particular catalytic compositepossesses unusual stability notwithstanding the presence of relativelylarge quantities of sulfurous compounds in the fresh charge stock. Thecatalytic composite comprises an alumina-containing refractory inorganicoxide, a molybdenum component, a platinum or palladium component and analkali-metal component, the latter being preferably potassium and/orlithium. It is especially preferred, for use in this particularhydrocarbon hydroprocessing scheme, that the catalytic composite besubstantially free from any "acid-acting" propensities. The catalyticcomposite, utilized in the second reaction zone for the primary purposeof effecting the destructive conversion of sulfurous compounds intohydrogen sulfide and hydrocarbons, is a composite of analumina-containing refractory inorganic oxide, a platinum or palladiumcomponent, and a molybdenum component. Through the utilization of aparticular sequence of processing steps, and the use of the foregoingdescribed catalytic composites, the formation of high molecular weightpolymers and copolymers is inhibited to a degree which permitsprocessing for an extended period of time. Briefly, this is accomplishedby initiating the hydrorefining reactions at temperatures below about500° F., at which temperatures the coke-forming reactions are notpromoted. The operating conditions within the second reaction zone aresuch that the sulfurous compounds are removed without incurring thedetrimental polymerization reactions otherwise resulting were it not forthe saturation of the conjugated di-olefinic hydrocarbons within thefirst reaction zone.

The hydrocarbon distillate charge stock, for example, a light naphthaby-product from a commercial cracking unit designed and operated for theproduction of ethylene, having a gravity of about 34.0° API, a brominenumber of about 35.0, a diene value of about 17.5 and containing about1,600 ppm. by weight of sulfur and 75.9 vol.% aromatic hydrocarbons, isadmixed with recycled hydrogen. This light naphtha co-product has aninitial boiling point of about 164° F. and an end boiling point of about333° F. The hydrogen circulation rate is within the range of from about1,000 to about 10,000 scf./Bbl., and preferably in the narrower range offrom 1,500 to about 6,000 scf./Bbl. The charge stock is heated to atemperature such that the maximum catalyst temperature is in the rangeof from about 200° F. to about 500° F., by way of heat-exchange withvarious product effluent streams, and is introduced into the firstreaction zone at an LHSV in the range of about 0.5 to about 10.0. Thereaction zone is maintained at a pressure of from 400 to about 1,000psig., and preferably at a level in the range of from 500 psig. to about900 psig.

The temperature of the product effluent from the first reaction zone isincreased to a level above about 500° F., and preferably to result in amaximum catalyst temperature in the range of 600° to 900° F. When theprocess is functioning efficiently, the diene value of the liquid chargeentering the second catalytic reaction zone is less than about 1.0 andoften less than about 0.3. The conversion of nitrogenous and sulfurouscompounds, and the saturation of mono-olefins, contained within thefirst zone effluent, is effected in the second zone. The secondcatalytic reaction zone is maintained under an imposed pressure of fromabout 400 to about 1,000 psig., and preferably at a level of from about500 to about 900 psig. The twostage process is facilitated when thefocal point for pressure control is the high pressure separator servingto separate the product effluent from the second catalytic reactionzone. It will, therefore, be maintained at a pressure slightly less thanthe first catalytic reaction zone, as a result of fluid flow through thesystem. The LHSV through the second reaction zone is about 0.5 to about10.0, based upon fresh feed only. The hydrogen circulation rate will bein a range of from 1,000 to about 10,000 scf./Bbl., and preferably fromabout 1,000 to about 8,000 scf./Bbl. Series-flow through the entiresystem is facilitated when the recycle hydrogen is admixed with thefresh hydrocarbon charge stock. Make-up hydrogen, to supplant thatconsumed in the overall process, may be introduced from any suitableexternal source, but is preferably introduced into the system by way ofthe effluent line from the first catalytic reaction zone to the secondcatalytic reaction zone.

With respect to the naphtha boiling range portion of the producteffluent, to sulfur concentration is about 0.1 ppm., the aromaticconcentration is about 75.1% by volume, the bromine number is less thanabout 0.3 and the diene value is essentially "nil."

With charge stocks having exceedingly high diene value, a recyclediluent is employed in order to present an excessive temperature rise inthe reaction system. Where so utilized, the source of the diluent ispreferably a portion of the normally liquid product effluent from thesecond catalytic reaction zone. The precise quantity of recycle materialvaries from feed stock to feed stock; however, the rate at any giventime is controlled by monitoring the diene value of the combined liquidfeed to the first reaction zone. As the diene value exceeds a level ofabout 25.0, the quantity of recycle is increased, thereby increasing thecombined liquid feed ratio; when the diene value approaches a level ofabout 20.0, or less, the quantity of recycle diluent may be lessened,thereby decreasing the combined liquid feed ratio.

With another so-called pyrolysis gasoline, haviang a gravity of about36.4° API, containing 600 ppm. by weight of sulfur, b 78.5% by volume ofaromatics, and having a bromine number of 45 and a diene value of 25.5it is initially processed in a first reaction zone containing acatalytic composite of alumina, 0.5% by weight of lithium, 0.20% byweight of palladium and 4% by weight of molybdenum, calculated as theelements. The fresh feed charge rate is 33,000 Sbl./day, and this isadmixed with 2,475 Bbl./day of the normally liquid diluent. Based uponfresh feed only, the LHSV is 2.5 and the hydrogen circulation rate is1,750 scf./Bbl. The charge is raised to a temperature of about 250° F.,and enters the first reaction zone at a pressure of about 840 psig. Theproduct effluent emanates from the first reaction zone at a pressure ofabout 830 psig. and a temperature of about 350° F. The effluent isadmixed with about 660 scf./Ebl. of make-up hydrogen, and thetemperature is increased to a level of about 545° F., the heated streamis introduced into the second reaction zone under a pressure of about790 psig. The LHSV, exclusive of the recycle diluent, is 2.5, and thehydrogen circulation rate is about 1,500. The second reaction zonecontains a catalyst of a composite of alumina, 0.375% by weight ofplatinum, and 4.0% by weight of molybdenum. The reaction producteffluent is introduced, following its use as a heat-exchange medium andfurther cooling to reduce its temperature from 620° to a level of 100°F., into a high-pressure separator at a pressure of about 750 psig. Thenormally liquid stream from the cold separator is introduced into areboiled stripping column for hydrogen sulfide removal anddepentanization. The hydrogen sulfide stripping column functions asconditions of temperature and pressure required to concentrate a C₆ toC₉ aromatic stream as a bottoms fraction. With respect to the overallproduct distribution, only 690 lbs./hr. of pentanes and lighterhydrocarbons is indicated in the stripper overhead. The aromaticconcentrate is recovered in an amount of about 40,070 lbs./hr. (thefresh feed is 40,120 lbs./hr.); these results are achieved with ahydrogen consumption of only 660 scf./Bbl. With respect to the desiredproduct, the aromatic concentration is 78.0, the sulfur concentration isless than 1.0 ppm. by weight, and the diene value is essentially "nil".

EXAMPLE III

This example is presented to illustrate still another hydrocarbonhydroprocessing scheme for the improvement of the jet fuelcharacteristics of sulfurous kerosene boiling range fractions. Theimprovement is especially noticeable in the IPT Smoke Point, theconcentration of aromatic hydrocarbons and the concentration of sulfur.A two-stage process wherein desulfurization is effected in the firstreaction zone at relatively mild severities which result in a normallyliquid product effluent containing from about 15 to about 35 ppm. ofsulfur by weight. Aromatic saturation is the principal reaction effectedin the second reaction zone, having disposed therein a catalyticcomposite of alumina, a halogen component, a platinum or palladiumcomponent, and a molybdenum component.

Suitable charge stocks are kerosene fractions having an initial boilingpoint as low as about 300° F., and an end boiling point as high as about575° F., and, in some instances, up to 600° F. Examplary of suchkerosene fractions are those boiling from about 300° to about 550° F.,from 320° to about 500° F., from 330° to about 530° F., etc. As aspecific example, a kerosene obtained from hydrocracking a Mid-continentslurry oil, having a gravity of about 30.5° API, an initial boilingpoint of about 388° F., an end boiling point of about 522° F., has anIPT Smoke Point of 17.1 mm., and contains 530 ppm. of sulfur and 34.8%by volume of aromatic hydrocarbons. Through the use of the catalyticprocess of the present invention, the improvement in the jet fuelquality of such a kerosene fraction is most significant with respect toraising the IPT Smoke Point, and reducing the concentration of sulfurand the quantity of aromatic hydrocarbons. Specifications regarding thepoorest quality of jet fuel, commonly referred to as Jet-A, Jet-Al andJet-B call for a sulfur concentration of about 0.3% by weight maximum(3,000 ppm.), a minimum IPT Smoke Point of 25 mm. and a maximum aromaticconcentration of about 20.0 vol.%.

The charge stock is admixed with circulating hydrogen in an amountwithin the range of from about 1,000 to about 2,000 scf./Bbl. Thismixture is heated to a temperature level necessary to control themaximum catalyst bed temperature below about 750° F., and preferably notabove 700° F., with a lower catalyst bed temperature of about 600° F.The catalyst, a desulfurization catalyst containing about 2.2% by weightof cobalt and about 5.7% by weight of molybdenum, composited withalumina is disposed in a reaction zone maintained under an imposedpressure in the range of from about 500 to about 1,000 psig. The LHSV isin the range of about 0.5 to about 10.0, and preferably from about 0.5to about 5.0. The total product effluent from this first catalyticreaction zone is separated to provide a hydrogen-rich gaseous phase anda normally liquid hydrocarbon stream containing 15 ppm. to about 35 ppm.of sulfur by weight. The normally liquid phase portion of the firstreaction zone effluent is utilized as the fresh feed charge stock to thesecond reaction zone. In this particular instance, the first reactionzone decreases the sulfur concentration to about 25 ppm., the aromaticconcentration to about 16.3% by volume, and has increased the IPT SmokePoint to a level of about 21.5 mm.

The catalytic composite within the second reaction zone comprisesalumina, 0.375% by weight of platinum, 4% by weight of molybdenum andabout 0.60% by weight of combined chloride, calculated on the basis ofthe elements. The reaction zone is maintained at a pressure of about 400to about 1,500 psig., and the hydrogen circulation rate is in the rangeof 1,500 to about 10,000 scf./Bbl. The LHSV, hereinbefore defined, is inthe range of from about 0.5 to about 5.0, and preferably from about 0.5to about 3.0. It is preferred to limit the catalyst bed temperature inthe second reaction zone to a maximum level of about 750° F. With acatalyst of this particular chemical and physical characteristics,optimum aromatic saturation, processing a feed stock containing fromabout 15 to about 35 ppm. of sulfur, is effected at maximum catalyst bedtemperature in the range of about 625° to about 750° F. With respect tothe normally liquid kerosene fraction, recovered from the condensedliquid removed from the total product effluent, the sulfur concentrationis effectively "nil," being about 0.1 ppm. The quantity of aromatichydrocarbons has been decreased to a level of about 0.75% by volume, andthe IPT Smoke Point has been increased to about 36.3 mm.

With respect to another kerosene fraction, having an IPT Smoke Point ofabout 20.3 mm., an aromatic concentration of about 19.3 vol.% and asulfur concentration of about 17 ppm. by weight, the same is processedin a catalytic reaction zone at a pressure of about 850 psig. and amaximum catalyst bed temperature of about 725° F. The LBSV is about1.35, and the hydrogen circulation rate is about 8,000 scf./Bbl. Thecatalytic composite disposed within the reaction zone comprises alumina,0.25% by weight of platinum, 4.0% by weight of molybdenum, about 0.35%by weight of combined chloride and 0.35% by weight of combined fluoride.Following separation and distillation, to concentrate the kerosenefraction, analyses indicate that the Smoke Point has been increased to alevel of about 36.9 mm., the aromatic concentration has been lowered toabout 0.6% by volume and the sulfur concentration is essentially "nil."

EXAMPLE IV

This illustration of a hydrocarbon hydroprocessing scheme, encompassedby my invention, is one which involves hydrocracking heavyhydrocarbonaceous material into lower-boiling hydrocarbon products. Inthis instance, the preferred catalysts contain a cobalt component, and amolybdenum component, combined with a crystalline aluminosilicatecarrier material, preferably faujasite, and still more preferably onewhich is at least 90.0% by weight zeolitic.

Most of the virgin stocks, intended for hydrocracking, are contaminantedby sulfurous compounds and nitrogenous compounds, and, in the case ofthe heavier charge stocks, various metallic contaminants, insolubleasphalts, etc. Contaminated charge stocks are generally hydrorefined inorder to prepare a charge suitable for hydrocracking. Thus, thecatalytic process of the present invention can be beneficially utilizedas the second stage of a two-stage process, in the first stage of whichthe fresh feed is hydrorefined.

Hydrocracking reactions are generally effected at elevated pressures inthe range of about 800 to about 5,000 psig., and preferably at someintermediate level of 1,000 to about 3,500 psig. Liquid hourly spacevelocities of about 0.25 to about 10.0 will be suitable, the lower rangegenerally reserved for the heavier stocks. The hydrogen circulation ratewill be at least about 3,000 scf./Bbl., with an upper limit of about50,000 scf./Bbl., based upon fresh feed. For the majority of feedstocks, hydrogen circulation in the range of 5,000 to 20,000 scf./Bbl.will suffice. With respect to the LBSV, it is based upon fresh feed,notwithstanding the use of recycle liquid providing a combined liquidfeed ratio in the range of about 1.25 to about 6.0. The operatingtemperature again alludes to the temperature of the catalyst within thereaction zone, and is in the range of about 400° to about 900° F. Sincethe principal reactions are exothermic in nature, the increasingtemperature gradient, experienced as the charge stock traverses thecatalyst bed, results in an outlet temperature higher than that at theinlet to the catalyst bed. The maximum catalyst temperature should notexceed 900° F., and it is generally a preferred technique to limit thetemperature increase to 100° F. or less.

Although amorphous composites of alumina and silica, containing fromabout 10.0% to about 90.0% by weight of the latter, are suitable for usein the catalytic composite employed in the present process, a preferredcarrier material constitutes a crystalline aluminosilicate, preferablyfaujasite, of which at least about 90.0% by weight is zeolitic. Thiscarrier material, and a method of preparing the same, have hereinbeforebeen described. A possible constituent of the catalyst is a halogencomponent, either fluorine, chlorine, iodine, bromine, or mixturesthereof. Of these, it is preferred to utilize a catalyst containingfluorine and/or chlorine. The halogen component will be composited withthe carrier material in such a manner as results in a final compositecontaining about 0.1% to about 1.5% by weight of halogen, calculated onan elemental basis.

A specific illustration of this hydrocarbon hydroprocessing techniqueinvolves the use of a catalytic composite of about 4.0% by weight ofmolybdenum, 4.0% by weight of cobalt and 0.7% by weight of combinedchlorine, combined with a crystalline aluminosilicate material of whichabout 90.9% by weight constitutes faujasite. This catalyst is intendedfor utilization in the conversion of 16,000 Bbl./day of a blend of lightgas oils to produce maximum quantities of a heptane-400° F. gasolineboiling range fraction. The charge stock has a gravity of 33.8° API,contains 0.19% by weight of sulfur (1,900 ppm.) and 67 ppm. by weight ofnitrogen, and has an initial boiling point of 639° F., a 50% volumetricdistillation temperature of 494° F. and an end boiling point of 655° F.The charge stock is initially subjected to a clean-up operation atmaximum catalyst temperature of 750° F., a combined feed ratio of 1.0,an LBSV of 2.41 with a hydrogen circulation rate of about 5,000scf./Bbl. The pressure imposed upon the catalyst within the reactionzone is about 1,500 psig. Since at least a portion of the blended gasoil charge stock will be converted into lower-boiling hydrocarbonproducts, the effluent from this clean-up reaction zone is separated toprovide a normally liquid, 400° F.-plus charge for the hydrocrackingreaction zone containing the catalyst. The pressure imposed upon thesecond reaction zone is about 1,300 psig., and the hydrogen circulationrate is about 8,000 scf./Bbl. The original quantity of fresh feed to theclean-up reaction zone is about 16,000 Bbl./day; following separation ofthe first zone effluent to provide the 400° F.-plus charge to the secondreaction zone, the charge to the second reaction zone is in an amount ofabout 12,150 Bbl./day, providing a LBSV of 0.85. The temperature at theinlet to the catalyst bed is 665° F., and a conventional hydrogen quenchstream is utilized to maintain the maximum reactor outlet temperature atabout 700° F. Following separation of the product effluent from thesecond reaction zone, to concentrate the desired gasoline boiling rangefraction, the remaining 400° F.-plus normally liquid material, in anamount of 7,290 Bbl./day, is recycled to the inlet of the secondreaction zone, thus providing a combined liquid feed ratio thereto ofabout 1.60. In the following table, there is indicated the product yieldand distribution of this process. With respect to normally liquidhydrocarbons, for convenience including butanes, the yields are given invol.%; with respect to the normally gaseous hydrocarbons, ammonia andhydrogen sulfide, the yields are given in terms of weight percent. Withrespect to the first reaction zone, the hydrogen consumption is 1.31% byweight of the fresh feed (741 scf./Bbl.), and for the hydrocrackingreaction zone, 1.26% by weight of the fresh feed charge stock, or 713scf./Bbl.

                  TABLE                                                           ______________________________________                                        Hydrocracking Product Yield and Distribution                                  Component      Stage I   Stage II    Total                                    ______________________________________                                        Ammonia        0.01      --          0.01                                     Hydrogen Sulfide                                                                             0.21      --          0.21                                     Methane        0.12      0.02        0.14                                     Ethane         0.22      0.40        0.62                                     Propane        1.03      3.48        4.51                                     Butanes        3.90      14.66       18.56                                    Pentanes       3.04      11.28       14.32                                    Hexanes        3.00      11.21       14.21                                    C.sub.7 -400° F.                                                                      18.85     49.56       68.41                                    400° F.-plus                                                                          75.92*    --          --                                       ______________________________________                                         *Charge to Stage II                                                      

With respect to both the butane product and pentane product, the formeris indicated as being about 68.0% isobutanes, while the latterconstitutes about 93.0% isopentanes. An analysis of the combinedpentane/hexane fraction indicates a gravity of 82.6° API, a clearresearch octane rating of 85.0 and a leaded research octane rating of99.0; it will be noted that this constitutes an excellent blendingcomponent for motor fuel. The desired heptane-400° F. product indicatesa gravity of 48.8° API, a clear research octane rating of 72.0 and aleaded research octane rating of 88.0. This gasoline boiling rangefraction constitutes about 34.0% by volume paraffins, 36.0% by volumenaphthenes and 30.0% by volume aromatic hydrocarbons. It will berecognized that this gasoline boiling range fraction constitutes anexcellent charge stock for a catalytic reforming unit to improve themotor fuel characteristics thereof.

The following comparative examples, including an example of onepreferred embodiment of the present invention, are presented inillustration of the improvement resulting from the method of manufactureof this invention and are not intended as an undue limitation on thegenerally broad scope of the invention as set out in the appendedclaims.

EXAMPLE V

About 450 grams of a commercial powdered alpha-alumina monohydrate(Catapal S) was thoroughly dry mixed with 95.6 grams of a finelypowdered, volailte free, molybdic oxide and about 19.9 grams of powderedcobalt carbonate. Approximately 245 grams of 13 weight percent nitricacid was then added to the powdered mixture in a muller, the mixturebeing thereby converted to a dough. The mixture was mulled for about anhour and thereafter extruded through a perforated plate comprising 1/32inch perforations. The extrudate was dried and calcined in air for aboutan hour at 750° F. and thereafter for an additional hour at 1100° F. Theextruded particles, broken to an average length of about 1/8 inch,contained 2.8 weight percent Co and 8.7 weight percent Mo.

EXAMPLE VI

Pursuant to the present invention, the extrudate particles of Example Iwere further impregnated with molybdic acid and cobalt nitrate. Thus,about 100 grams of the extrudate particles were impregnated with acommon ammoniscal solution of molybdic acid and cobalt nitrate preparedby commingling an aqueous solution of 2.7 grams of 85% molybdic acid and2.3 milliliters of ammonium hydroxide with an aqueous solution of 1.2grams of cobalt nitrate hexahydrate and 1.2 milliliters of ammoniumhydroxide, the resulting solution being subsequently diluted to about170 milliliters with water. The extrudate particles were immersed in theimpregnating solution which was then evaporated to dryness. Theimpregnated particles were then calcined in air for about 1 hour at 750°F. and for an additional hour at 1100° F. The extrudate particlescontained 3.5 weight percent Co and 10.3 weight percent Mo.

EXAMPLE VII

In this example, the cobalt and molybdenum components were incorporatedin the catalytic composite solely by impregnation. In this example, 100grams of the powdered alpha-alumina monohydrate was mulled with about 55grams of 13 weight percent nitric acid to form a dough. The dough wasthen extruded, dried, and calcined in air for about an hour at 750° F.and then for an additional hour at 1100° F. The calcined particles wereimmersed in a common ammoniacal solution of molybdic acid and cobaltnitrate hexahydrate prepared by commingling an aqueous solution of 20.7grams of 85% molybdic acid and 12 milliliters of ammonium hydroxide withan aqueous solution of 16 grams of cobalt nitrate hexahydrate and 12milliliters of ammonium hydroxide. Approximately 87 grams of theextrudate particles were immersed in the impregnating solution which wasthen evaporated to dryness. The impregnated particles were then calcinedas heretofore described. The impregnated extrudate particles contained3.25 weight percent Co and 9.4 weight percent Mo.

A summary of catalyst properties and activity test results is tabulatedbelow.

    ______________________________________                                        Promoter Addition      Coextrusion &                                                                            Impreg-                                       Technique Coextrusion                                                                              Impregnation                                                                             nation                                      ______________________________________                                        Catalyst Properties                                                            ABD        0.685      0.733      0.697                                        Piece Density,                                                                g/cc       1.25       1.30       1.28                                         Diameter, In.                                                                            0.028      0.028      0.029                                        Wt. % Co   2.8        3.5        3.25                                         Wt. % Mo   8.7        10.3       9.4                                          SA, m.sup.2 /g                                                                           299        294        272                                          PV, cc/g   0.51       0.48       0.51                                         PD, A      68         65         75                                          ______________________________________                                    

The above-described catalysts were evaluated with respect to thedesulfurization of a vacuum gas oil boiling in the 600°-1050° F. rangeand containing 2.6 weight percent sulfur. In each case, the catalyst wasdisposed as a fixed bed in a vertical tubular reactor maintained at 650psig. and 750° F. The vacuum gas oil was charged over the catalyst at3.0 liquid hourly space velocity in admixture with 1800 standard cubicfeet of hydrogen per barrel of charge. The reactor effluent wasseparated into a liquid and a gaseous phase in a high pressure separatorat 250° F., and the liquid phase was treated in a stripper column forthe separation of light ends. The liquid stripper bottoms collected overan 8-hour period was analyzed for sulfur.

In the described desulfurization of vacuum gas oil, the catalyst ofExample VI, wherein the cobalt and molybdenum components wereincorporated in the catalyst by coextrusion followed by impregnationpursuant to the method of this invention, was 55% more active than thecatalyst of Example VII, wherein the cobalt and molybdenum componentswere incorporated solely by impregnation, and 95% more active than thecatalyst of Example V, wherein the cobalt and molybdenum components wereincorporated solely by coextrusion with the alumina.

The foregoing specification, and particularly the examples, indicatesthe method by which the present invention is effected, and the benefitsafforded through the utilization thereof.

I claim as my invention:
 1. A process for desulfurizing a sulfuroushydrocarbon distillate containing mono-olefinic hydrocarbons andaromatics hydrocarbons which process comprises reacting said distillatewith hydrogen in contact with a catalytic composite of from about 4 toabout 30 weight percent Group VI-B metal, from about 0.1 to about 30percent Group VIII metal and a refractory inorganic oxide carriermaterial which is prepared in the following manner:(a) dry mixing afinely divided Group VI-B metal compound, a Group VIII metal compoundand a refractory inorganic oxide, said metal compound being utilized inan amount to provide from about 25 to about 90% of the Group VI-B metalcomponent of the final catalytic composite, and from about 25 to about90% of the Group VIII metal component; (b) peptizing the mixture andforming an extrudable dough; (c) extruding said dough and calcining theextrudate; (d) impregnating the calcined extrudate with a Group VI-Bmetal compound and a Group VIII metal compound, said metal compoundbeing utilized in an amount to provide a final catalytic compositecontaining from about 4 to about 30 weight percent Group VI-B metal andfrom about 0.1 to about 10 weight percent Group VIII metal; and, (e)calcining the resulting composite in an oxidizing atmosphere.
 2. Aprocess for hydrogenating a coke-forming hydrocarbon distillatecontaining di-olefinic, mono-olefinic and aromatic hydrocarbons tohydrogenate di-olefinic hydrocarbons to monoolefinic hydrocarbons whichprocess comprises reacting said distillate with hydrogen at atemperature below about 500° F. in contact with a catalytic composite offrom about 4 to about 30 weight percent Group VI-B metal, from about 0.1to about 30 weight percent Group VIII metal and a refractory inorganicoxide carrier material which is prepared in the following manner:(a) drymixing a finely divided Group VI-B compound, a Group VIII metal compoundand a refractory inorganic oxide, said metal compound being utilized inan amount to provide from about 25 to about 90% of the Group VI-B metalcomponent of the final catalytic composite, and from about 25 to about90% of the Group VIII metal component; (b) peptizing the mixture andforming an extrudable dough; (c) extruding said dough and calcining theextrudate; (d) impregnating the calcined extrudate with a Group VI-Bmetal compound and a Group VIII metal compound, said metal compoundbeing utilized in an amount to provide a final catalytic compositecontaining from about 4 to about 30 weight percent Group VI-B metal andfrom about 0.1 to about 10 weight percent Group VIII metal; and, (e)calcining the resulting composite in an oxidizing atmosphere.
 3. Aprocess for improving the jet fuel characteristics of a sulfurouskerosene boiling range fraction by reducing the sulfur content of saidfraction which process comprises reacting said kerosene fraction andhydrogen at a maximum catalyst temperature of about 750° F. in acatalytic reaction zone containing a catalytic composite of from about 4to about 30 weight percent Group VI-B metal, from about 0.1 to about 30weight percent Group VIII metal and a refractory inorganic oxide carriermaterial which is prepared in the following manner:(a) dry mixing afinely divided Group VI-B metal compound, a Group VIII metal compoundand a refractory inorganic oxide, said metal compound being utilized inan amount to provide from about 25 to about 90% of the Group VI-B metalcomponent of the final catalytic composite, and from about 25 to about90% of the Group VIII metal component; (b) peptizing the mixture andforming an extrudable dough; (c) extruding said dough and calcining theextrudate; (d) impregnating the calcined extrudate with a Group VI-Bmetal compound and a Group VIII metal compound, said metal compoundbeing utilized in an amount to provide a final catalytic compositecontaining from about 4 to about 30 weight percent Group VI-B metal andfrom about 0.1 to about 10 weight percent Group VIII metal; and, (e)calcining the resulting composite in an oxidizing atmosphere.